Process for the production of gaseous olefins from petroleum distillate feedstocks

ABSTRACT

Olefins, e.g., ethylene and propylene, are produced by hydrogenating a petroleum distillate feedstock, e.g., a petroleum vacuum distillate feedstock (b. pt. 300* - 650*C), over a catalyst containing combinations of Ni, Mo, Co and Won SiO2/Al2O3 bases, preferably Ni/W/SiO2/Al2O3, under reaction conditions such that a substantial quantity of aromatics in the feedstock are hydrogenated. The whole product of hydrogenation is then thermally cracked in the presence of steam, producing an increased yield of gaseous olefin products.

United States Patent [1 1 Davis et al.

PROCESS FOR THE PRODUCTION OF GASEOUS OLEFINS FROM PETROLEUM DISTILLATE FEEDSTOCKS Inventors: Paul Trevor Davis, Feltham;

Terence George Glover, lsleworth; John- Robert Jones, Walton-on-Thames, all of England Assignee: BP Chemicals International Limited,

London, England Filed: Dec. 30, 1971 Appl. No.: 214,468

Foreign Application Priority Data Dec. 25, 1973 [5 6] References Cited UNITED STATES PATENTS 3,720,729 3/1973 Sze et a1. 3,513,217 5/1970 Raymond... 3,598,720 8/1971 Stolfa 208/57 3,306,844 2/1967 Brierley et al. 260/683 2,093,843 9/1937 McKee 208/57 Primary Examiner-l-lerbert Levine Attorney-Jacobs & Jacobs [57] ABSTRACT Olefins, e.g., ethylene and propylene, are produced by hydrogenating a petroleum distillate feedstock, e.g., a petroleum vacuum distillate feedstock (b. pt. 300 650C), over a catalyst containing combinations of Ni, Mo, Co and Won SiO /Al O bases, preferably Ni/W- /SiO /Al 0 under reaction conditions such that a substantial quantity of aromatics in the feedstock are hydrogenated. The whole product of hydrogenation is then thermally cracked in the presence of steam, producing an increased yield of gaseous olefin products.

6 Claims, No Drawings PROCESS FOR THE PRODUCTEON OF GASEOUS OLEFINS FROM PETROLEUM DHSTTLLATE FEEDSTOCKS The present invention relates to a process for the production of gaseous olefins from petroleum distillate feedstocks.

Ethylene, propylene and butadiene which are basic intermediates for a large proportion of the petrochemical industry are obtained in the main by thermal cracking of petroleum gases and distillates such as naphtha and gas oil. There is a world wide increase in demand on the use of these lighter components of petroleum and it is desirable that heavier feedstocks be utilised for olefin production. In the past a number of problems arose in the cracking of heavier feedstocks which have so far prevented their use in the economic production of light olefins. The principle problems were:

1. Excessive coke deposition in the cracking tubes which reduces heat transfer, thereby necessitating higher tube skin temperatures. Excessive coke deposition also restricts flow in the cracking tubes, and may ultimately lead to blockage. The coke must be removed at frequent intervals by burning out, involving plant shut-down for an excessive proportion of time onstream.

2. Tar deposition in the transfer lines and heat exchangers; this reduces the efficiency of heat recovery and requires plant shut-down for cleaning, again impairing the overall efficiency of operation.

3. Yields of olefin products are low compared with those from lighter feedstocks necessitating increased feedstock and fuel requirements with extra furnaces heat exchangers and other equipment involving much higher initial capital investment.

4. Feedstocks from many sources contain high levels of sulphur; this is not necessarily detrimental to the operation of the cracking process but may increase the cost of plant construction. Further the bulk of the sulphur is concentrated in the liquid products boiling above 200C which are therefore less valuable as fuel oil.

We have now found that by pretreating the petroleum distillate feedstock with hydrogen in the presence of a catalyst under conditions which lead to substantial reduction in the content of aromatic compounds (especially of polycyclic aromatic compounds) and of sulphur compounds, but without substantial reorganisation or breakdown of the carbon structures of the various component compounds, the disadvantages of the prior art referred to above are substantially overcome.

Thus according to the present invention there is provided a process for the production of olefins which process comprises hydrogenating a petroleum distillate feedstock in the presence of a hydrogenation catalyst and hydrogen and thermally cracking the resulting hydrogenated product in the presence of steam.

Thermal cracking within the context of this application is intended to include steam cracking but not catalytic cracking.

The preferred petroleum distillate feedstock is a vacuum distillate boiling within the range (at atmospheric pressure) 300 to 650C, though lighter distillate feedstocks such as gas oil boiling within the range 200 350C may be used.

' [t is important to avoid excessive breakdown of the feedstock in a hydrocracking type of reaction. A limited amount of breakdown can be tolerated and may even give the benefit of producing a more mobile product but excessive hydrocracking leads to the use of larger quantities of hydrogen with increased manufacturing costs and to the formation of products which do not give corresponding benefits in further increases in the yield of olefins. Any catalyst which is capable of catalysing the hydrogenation of compounds containing aromatic rings without substantial structural alteration or breakdwon may be used. Since most feedstocks contain sulphur and nitrogen compounds it is desirable that the catalyst should also possess some tolerance to these materials and their hydrogenation products. Hydrogenation catalysts embodying these requisites include for example nickel/molybdenum/alumina, cobalt/tungsten/alumina, nickel/tungsten/alumina, cobalt/molybdenum/alumina, nickel/cobalt/molybdenumlalumina, cobalt/mo]ybdenum/silica/alumina, nickel/molybdenum/silica/alumina, cobalt/tungsten/silica/alumina. A particularly active hydrogenation catalyst is nickel/- tungsten/silica/alumina.

Although it will usually be convenient to employ the hydrogenation catalyst without prior exposure to materials containing sulfur at least initially, the catalyst may also be used in the sulfided form.

The catalysts may conveniently be prepared by impregnating the support with an aqueous solution of a salt of each of the metals, either consecutively or simultaneously. Thus nickel may be added in the form of nickel nitrate, tungsten as ammonium metatungstate, cobalt as cobalt nitrate, acetate, etc. and molybdenum as ammonium molybdate. It will usually be found convenient to impregnate the support first with the salt of the metal which is to be present in the highest concentration in the finished catalyst, though this is not essential. Other methods of preparing the catalyst include precipitating the metals on the support from a solution of their salts and coprecipitation of the metals with the hydrated support material.

It is preferred that the catalysts be activated before use in the reaction by contact with a stream of hydrogen at a temperature in the range to 800C, preferably 300 to 600C, for a period of 1 minute to 24 hours. The sulfided form of the catalyst may conveniently be prepared by passing hydrogen through liquid tetrahydrothiophene and then over the catalyst maintained at a temperature in the range 100C to 800C, preferably 300C to 600C, for a period of 1 minute to 24 hours. I

Whilst the precise nature of the active species in the above hydrogenation catalysts is not known it is possible that the catalyst contains, in addition to the support, elemental metal, metal oxides, metal sulfides and complex aluminium or silicon/metal compounds.

Using nickel and cobalt catalysts the hydrogenation temperature may be in the range 50 to 500C, preferably 300 to 400C, and the pressure may be in the range 50 to 5000 p.s.i.g., preferably 200 to 3000 p.s.i.g.

The hydrocarbon Liquid Hourly Space Velocity (LHSV) may be in the range of 0.1 to 5.0, preferably 0.25 to 2.0.

Hydrogen is preferably used on-a recycle basis, preferably at about 5 to 10 times the molar rate of the hydrocarbon feedstock, and may be passed through scrubbers to remove hydrogen sulfide and ammonia before recycle. However other methods of operation may also be used such as batch operation in an autoclave.

For catalysts other than those containing cobalt or nickel the reaction conditions may be different.

Hydrogenation may be carried out in a single stage or in a series of two or more operations using the same or different catalysts.

The feedstock from the hydrogenation reaction is vaporised in the presence of steam at a steam to hydrocarbon weight ratio of about 0.521 to 20:1 and passed through a heated zone, preferably a tube, at a maximum temperature in the range 700 to 1,000C with a residence time in this temperature range between 0.01 and 5 seconds, preferably 0.1 to 2 seconds. The products are rapidly cooled in a heat exchange system and separated and purified by conventional means.

The invention is illustrated by the following examples:

COMPARISON TEST 1 A full range Agha Jari vacuum distillate with a hydrogen to carbon atomic ratio of 1.73 and a sulfur content of 1.72 per cent weight was steam cracked in a 26 ml quartz reactor at a maximum temperature of 830C. Analysis by physical separation and spectroscopic methods (including U.V. Absorbance) indicated that the aromatic compound content was 49% weight.

The steam to hydrocarbon feed weight ratio was 1 to 1 with an average total molar flow of 3.3 moles per hour. The ethylene and propylene yields were 23 and per cent weight respectively with a total conversion to cracked gas of 53 per cent weight on feed. Coke deposited in the cracking zone corresponded to 1,200 ppm of the hydrocarbon feed.

This example is provided for purposes of comparison and is not an example according to the invention.

EXAMPLE 1 The catalyst was prepared by calcining alumina at a temperature of 550C. The calcined alumina was then impregnated with an aqueous solution of ammonium molybdate, evaporated to dryness and further calcined at 550C. This procedure was then repeated using an aqueous solution of cobalt nitrate. The catalyst was then activated in a stream of hydrogen at 400C for 16 hours.

A 250g sample of the Agha Jari vacuum distillate used in Comparison Test 1 was hydrogenated in a 1 litre rocking autoclave at 350C under 1,500 p.s.i.g. of hydrogen during 8 h using 100g of the cobalt/molybdenum/alumina catalyst prepared as above. The recovered hydrogenated vacuum distillate, sample A, had a hydrogen to carbon atomic ratio of 1.84 and a sulfur content of less than 0.05 per cent weight. Analysis by physical separation and spectroscopic methods (including U.V. Absorbance) indicated that the aromatic compound content was 19% weight. This material was steam cracked under the same conditions as were used in Comparison Test 1. The ethylene and propylene yields were 26 and 10 per cent weight on feed respectively with a total conversion to cracked gas of 58 per cent. There was also a substantial reduction in the heavier products compared with those formed from untreated vacuum distillate. In particular tarry material condensed in the transfer line from the reactor was formed in only half the amount observed in Comparison Test 1. The amount of coke deposited in the reactor zone was only 250 ppm on feed.

EXAMPLE 2 A g sample of the hydrogenated vacuum distillate sample A, was further hydrogenated in a rocking autoclave at 350C and 1,500 psig of hydrogen during 18 hours using 40 g of a 5 per cent nickel on silica catalyst prepared by impregnation as in Example 1. This further hydrogenated vacuum distillate has a hydrogen to carbon atomic ratio of 1.91 and a sulphur content of less than 0.02 per cent weight. Analysis by physical separation and spectroscopic methods (including U.V. Absorbance) indicated that the aromatic compound content was less than 2 per cent weight. On steam cracking this sample under the conditions used in Comparison Test 1 the ethylene and propylene yields were found to be 29 and 1 1 per cent weight respectively, and the total cracked gas yield has increased to 64 per cent. The yields of fuel oil and tarry material were reduced to a quarter of the values obtained from the untreated vacuum distillate while the amount of coke deposited in the reactor was only 100 ppm on feed.

COMPARISON TEST 2 A full range Kuwait vacuum distillate with a hydrogen to carbon atomic ratio of 1.74 and a sulfur content of 2.78 per cent weight was steam cracked in a 20 ml quartz reactor at a temperature of 830C. Analysis by physical separation and spectroscopic methods (including UV absorbance) indicated that the aromatic compound content was 52 per cent weight.

The steam to hydrocarbon feed weight ratio was 1 to 1 with an average hydrocarbon feed rate of 62 g per hour. The ethylene and propylene yields were 23 and 10 per cent weight respectively with 52 per cent weight of the feed converted to cracked gas. Coke deposited in the cracking zone corresponded to 1,050 ppm weight of the hydrocarbon feed. The sulfur content of the fuel oil was 6.8 per cent weight.

This Example is provided for purposes of comparison and is not an example according to the invention.

EXAMPLE 3 A 200 g sample of Kuwait vacuum distillate used in Comparison Test 2 was hydrogenated in a 1 litre rocking autoclave at 350C under 2,200 psig of hydrogen during 8 hours using 50 g of cobalt/molybdenumlalumina catalyst prepared by impregnation as in Example 1.

The recovered hydrogenated vacuum distillate has a hydrogen to carbon atomic ratio of 1.87 and a sulfur content of less than 400 ppm weight. Analysis by physical separation and spectroscopic methods (including UV absorbance) indicated that the aromatic compound content was 16 per cent weight.

This material was steam cracked under the same conditions as Comparison Test 2. The ethylene and propylene yields were 27 and 12 per cent weight on feed respectively with 62 per cent weight of the feed converted to cracked gas. The yields of fuel oil and tarry material were reduced to one third of the values obtained from the untreated vacuum distillate. Coke deposited in the cracking zone corresponded to ppm weight of the hydrocarbon feed.

EXAMPLE 4 A 200 g sample of Kuwait vacuum distillate used in Comparison Test 2 was hydrogenated in a 1 litre rocking autoclave at 350C under 2200 psig of hydrogen during 8 hours using 50 g of nickel/tungsten/- silica/alumina catalyst prepared by impregnation as in Example 1. The recovered hydrogenated vacuum distillate had a hydrogen to carbon atomic ratio of 1.95 and a sulfur content of 100 ppm weight. Analysis by physical separation and spectroscopic methods (including UV absorbance) indicated that the aromatic compound content was less than 2 per cent weight.

This material was steam cracked under the same conditions as Comparison Test 2. The ethylene and propylene yields were 28 and 14 per cent weight on feed respectively with 68 per cent of the feed converted to cracked gas. The yields of fuel oil and tarry material were reduced to one quarter of the values obtained from the untreated vacuum distillate while the amount of coke deposited in the reactor was only 100 ppm. The sulfur content of the fuel oil was only 300 ppm.

EXAMPLE 5 A 200 g sample of the Kuwait vacuum distillate used in Comparison Test 2 was hydrogenated in a 3 litre rocking autoclave at 350C under 2200 psig of hydrogen during 75 hours using 60.0 g of a cobalt/molybdenum/alumina catalyst prepared by impregnation as in Example 1. The recovered hydrogenated vacuum distillate had a hydrogen to carbon atomic ratio of 1.95 and a sulphur content of 125 ppm weight. Analysis by physical separation and spectroscopic methods (including UV absorbance) indicated that the aromatic compound content was less than 2 per cent weight.

This material was steam cracked under the same conditions as Comparison Test 2. The ethylene and propylene yields were 28 and 14 per cent weight on feed respectively, with 67 per cent weight of the feed converted to cracked gas. The yields of fuel oil and tarry material were reduced to a quarter of the values obtained from the untreated vacuum distillate while the amount of coke deposited in the reactor was only 100 ppm.

EXAMPLE 6 A 200 g sample of Kuwait vacuum distillate used in Comparison Test 2 was hydrogenated in a 1 litre rocking autoclave at 350C under 2,200 psig of hydrogen during 8 hours using 50 g of nickel/tungsten/alumina catalyst prepared by impregnation as in Example 1. The recovered hydrogenated vaccum distillate had a hydrogen to carbon atomic ratio of 1.88 and a sulphur content of less than 300 ppm weight. Analysis by physical separation and spectrosopic methods (including UV absorbance) indicated that the aromatic compound content was per cent weight.

This material was steam cracked under the same conditions as Comparison Test 2. The ethylene and propylene yields were 27 and 13 per cent weight on feed respectively with 62 per cent weight of the feed converted to cracked gas. The yields of fuel oil and tarry material were reduced to one third of the values obtained from untreated vacuum distillate, while the amount ofcoke deposited in the reactor was only l20 pp Comparison Tests 1 and 2 show the ethylene and propylene yields, cracked gas yield and coke lay-down resulting from the cracking of straight run wax distillate without pretreatment. The Examples demonstrate that hydrogenation of wax distillates prior to cracking leads to an improvement in all the above parameters and, in addition, not only is the sulfur content reduced but the yields of fuel oil and tarry material are reduced also. The increased degree of hydrogenation achieved by using a nickel/tungsten/silica alumina catalyst is demonstrated in Example 4. The degree of hydrogenation is similar to that of Example 5 but is achieved in a much shorter time period.

In all examples according to the invention it was further observed that the fuel oil product had a much lower viscosity and a reduced tendency for emulsific'ation with water after condensation of the product and diluent steam by comparison with the fuel oil produced in the Comparison Tests, thus facilitating both pumping and handling at lower temperatures and separation and fractionation of the liquid products respectively.

We claim:

1. A process for the production of olefins which process comprises hydrogenating a petroleum vacuum distillate feedstock boiling in the range of 300 to 650C. in the presence of a nickel/tungsten/silica/alumina hydrogenation catalyst and hydrogen at a temperature in the range 50 to 500C and a pressure in the range 50 to 5,000 psig and thermally cracking the resulting hydrogenated product in the presence of steam.

2. A process according to claim 1 wherein the hydrogenation temperature is in the range 300 to 400C.

3. A process according to claim 1 wherein the hydrogenation pressure is in the range 200 to 3,000 psig.

v 4. A process according to claim 1 wherein the hydrocarbon liquid Hourly Space Velocity (L.H.S.V.) is in the range 0.1 to 5.0.

5. A process according to claim 1 wherein the hydrogen is used on a recycle basis at 5 to 10 times the molar rate of the hydrocarbon feedstock.

6. A process according to claim 1 wherein the feedstock from the hydrogenation reaction is varpourised in the presence of steam, at a steam to hydrocarbon weight ratio of 0.5:1 to 20:1 and is passed through a heated zone at a maximum temperature in the range 700 to l,000 with a residence time in this temperature range between 0.01 and 5 seconds. 

2. A process according to claim 1 wherein the hydrogenatioN temperature is in the range 300* to 400*C.
 3. A process according to claim 1 wherein the hydrogenation pressure is in the range 200 to 3,000 psig.
 4. A process according to claim 1 wherein the hydrocarbon liquid Hourly Space Velocity (L.H.S.V.) is in the range 0.1 to 5.0.
 5. A process according to claim 1 wherein the hydrogen is used on a recycle basis at 5 to 10 times the molar rate of the hydrocarbon feedstock.
 6. A process according to claim 1 wherein the feedstock from the hydrogenation reaction is varpourised in the presence of steam, at a steam to hydrocarbon weight ratio of 0.5:1 to 2.0:1 and is passed through a heated zone at a maximum temperature in the range 700* to 1,000* with a residence time in this temperature range between 0.01 and 5 seconds. 